ES2409604T3 - Procedure for the conversion of an alcohol or alcohols into an alcohol or alcohols with an increased carbon chain - Google Patents

Abstract

A procedure for the approval of an alcohol or alcohols, characterized by the following steps: 1. introducing an alcohol or alcohols in an etherification unit to produce an ether or alkyl ethers, 2. introducing at least a part of the ether or alkyl ethers from step 1, together with CO and optionally H2, in a carbonylation unit, in the presence of a homogeneous or heterogeneous carbonylation catalyst, to produce an alkyl ester or esters, 3. introduce at least a portion of the ester or alkyl esters from stage 2, together with H2, in a hydrogenation unit, to produce an approved alcohol or alcohols, 4. optionally recycle at least a part of the alcohol or the approved alcohols from stage 3 to the etherification unit of stage 1 , and 5. recovering an alcohol or alcohols from the hydrogenation unit of step 3.

Description

Procedure for the conversion of an alcohol or alcohols into an alcohol or alcohols with an increased carbon chain.

The present invention relates to a process for the conversion of an alcohol or alcohols into a higher alcohol or alcohols. The present invention also relates to a process for the production of alcohols from synthesis gas.

In particular, the present invention relates to the production of an alcohol or alcohols from a carbonaceous raw material; wherein the carbonaceous raw material is first converted into synthesis gas which is then converted into an alkyl carboxylate ester or esters which are then hydrogenated to produce an alcohol or alcohols.

In recent years, the increased use and demand of alcohols such as methanol, ethanol and higher alcohols has led to increased interest in procedures related to the production of alcohols. Such alcohols can be produced by fermentation of, for example, sugars and / or cellulosic materials.

Alternatively, alcohols, such as ethanol, can be produced from synthesis gas. Synthesis gas refers to a combination of H2 and carbon oxides produced in a synthesis gas plant from a carbon source such as natural gas, petroleum liquids, biomass and other carbonaceous materials that include coal, recycled plastics, waste municipal or any organic material. Thus, alcohol and alcohol derivatives can provide non-oil based routes for the production of valuable chemicals and fuels.

Generally, the production of alcohols, for example methanol, takes place through three processing steps: synthesis gas preparation, methanol synthesis and methanol purification. In the synthesis gas preparation stage, an additional phase in which the feedstock is treated can be used,

p. ex. The raw material is purified to remove sulfur and other potential catalyst poisons before becoming synthesis gas. This treatment can also be carried out after the preparation of the synthesis gas; for example, when biomass or coal is used.

The reaction to produce an alcohol or alcohols from synthesis gas is generally exothermic. It is believed that the formation of C2 and C2 + alcohols progresses through the formation of methanol for modified methanol catalysts and cobalt and molybdenum sulfide catalysts. However, methanol production is limited by equilibrium and thus requires high pressures in order to achieve viable yields. Hence, pressure can be used to increase the yield, since the reaction that produces methanol exhibits a decrease in volume, as disclosed in US 3326956.

A low pressure copper-based methanol synthesis catalyst is commercially available from suppliers such as BASF, Johnson Matthey and Haldor-Topsoe. The methanol yields from copper-based catalysts are generally above 99.5% of converted CO + CO2 present. Water is a byproduct of the conversion of CO2 into methanol and the conversion of CO, synthesis gas, into C2 and C2 + oxygenates. In the presence of an active water-gas displacement catalyst, such as a methanol catalyst or a cobalt-molybdenum catalyst, the water is equilibrated with CO to give CO2 and H2. A document entitled "Selection of Technology for Large Methanol Plants" by Helge Holm-Larsen, presented at the 1994 World Methanol Conference, Nov. 30-Dec. 1. 1994, in Geneva, Switzerland, reviews the advances in methanol production and shows how an additional reduction in methanol production costs will result in the construction of very large plants with capacities that reach 10,000 t per day.

Other procedures for the production of a C2 and C2 + alcohol or alcohols include the procedures described later herein.

WO 8303409 describes a process by which ethanol is produced by carbonylation of methanol by reaction with CO in the presence of a carbonylation catalyst to form ethanoic acid which is then converted into an ethanolate ester followed by hydrogenation of the ethanoate ester. formed to give ethanol or a mixture of ethanol and other alcohol that can be separated by distillation. The carbonylation can be carried out using a mixture of CO / H2 and the hydrogenolysis can be carried out similarly in the presence of CO, leading to the possibility of circulating gas between the carbonylation and hydrogenolysis areas with synthesis gas, preferably a molar mixture of H2: CO 2: 1 that is used as a composition gas.

US 4122110 refers to a process for manufacturing alcohols, particularly linear saturated primary alcohols, by reacting CO with H2 at a pressure between 2 and 25 MPa and a temperature between 150 and 400 ° C, in the presence of a catalyst, characterized in that the catalyst It contains at least 4 essential elements: (a)

copper (b) cobalt (c) at least one M element selected from chromium, iron, vanadium and manganese, and (d) at least one alkali metal.

US 4831060 refers to the production of mixed alcohols from gaseous CO and H2 using a catalyst, optionally with a cocatalyst, wherein the catalyst metals are molybdenum, tungsten or rhenium, and the cocatalyst metals are cobalt, nickel or iron. . The catalyst is promoted with a Fischer-Tropsch catalyst such as an alkaline or alkaline earth series metal or a smaller amount of thorium and is further treated by sulfurization. The composition of the mixed alcohol fraction can be selected by selecting the extent of intimate contact between the catalytic components.

Journal of Catalysis, 1988, 114, 90-99 discloses an ethanol formation mechanism from synthesis gas on CuO / ZnO / Al2O3. The formation of ethanol from CO and H2 on a catalyst for CuO / ZnO methanol is studied in a fixed bed microreactor by measuring the isotopic distribution of carbon in ethanol obtained as a product when isotopically enriched methanol 13C was added to the feed.

US 4,454,358 discloses a continuous process for the production of ethanol through the carbonylation of methanol, by (a) carbonylating methanol, in an R reactor, in the presence of a carbonyl complex of a group VIII metal of the periodic table and of a halogenated compound, (b) separating, in a distillation column D1, the discharge from the reactor, in an upper fraction comprising methyl acetate, methanol, dimethyl ether and an organohalogenated compound, and in a lower fraction comprising water, small amounts of acetic acid and the catalyst, if the latter is not in a fixed bed, adjusting the residence time so that most of the acetic acid reacts with the methanol present to give methyl acetate, (c) separating the fraction upper from D1, in a distillation column D2, in a higher fraction comprising small amounts of methyl acetate, methanol, dimethyl ether and the organohalogenated compound, and a lower fraction comprising methyl acetate and methanol, and recycling the upper fraction to the reactor R, (d) distilling, through the upper part of the distillation column D3, most of the water from the fraction lower from D1 and remove the water from the circulation, and recycle to the reactor R the lower fraction consisting of small amounts of water, acetic acid and the catalyst, (e) use hydrogen to hydrogenate, in the hydrogenation reactor H, the lower fraction from D2, in a conventional manner, to give a mixture of methanol and ethanol, and (f) separate the mixture into ethanol and methanol in a distillation column D4, and recycle the methanol to the reactor

R.

US 4,831,060 discloses a process for selectively producing mixed alcohols from synthesis gas comprising contacting a mixture of hydrogen and carbon monoxide with a catalytic amount of a catalyst wherein the catalyst is of components of

(one)

 a catalytically active metal of molybdenum, tungsten or rhenium, in free or combined form;

(2)

 a cobalt, nickel or iron cocatalytic metal, in free or combined form;

(3)

 a Fischer-Tropsch promoter; Y

(4)

 an optional support;

said components combined by dry mixing, mixing as a wet paste, wet impregnation or, if the first component is rhenium, coprecipitation, and then sulfurized, under conditions sufficient to form said product in at least 20 percent by weight selectivity CO2 free carbon.

WO 2007/117590 A2 discloses a method for producing ethanol from synthesis gas comprising the steps of providing synthesis gas; hydrogen; reacting a catalyst and a promoter with methanol, carbon monoxide and hydrogen to produce a mixture comprising methyl acetate, hydrogen, methanol, acetic acid and water; separating the mixture to separate the promoter, a mixture of methyl acetate and methanol, and a mixture of acetic acid and water; add hydrogen to the mixture of methyl acetate and methanol and react the mixture with a hydrogenation catalyst to produce ethanol. WO 2007/117590 A2 also discloses a process that converts synthesis gas into ethanol. The synthesis gas is converted into a mixture of methanol, carbon monoxide and hydrogen; the mixture of methanol and carbon monoxide is reacted using a catalyst to produce acetic acid and carbon monoxide, the acetic acid is reacted with ethanol in the presence of a catalyst to produce a mixture comprising ethyl acetate and water; and ethyl acetate is reacted with hydrogen using a hydrogenation catalyst to produce ethanol.

US 2006/252959 A1 discloses a process for producing a product comprising a lower alkyl ester of a lower aliphatic carboxylic acid, which comprises reacting a lower alkyl ether with carbon monoxide in the presence of a catalyst comprising mordenite and / or low ferrite substantially anhydrous conditions.

As the importance of alcohols is always growing in today's world, it is necessary and desired to produce alcohols with superior carbon efficiency, superior conversion and improved selectivity from a carbonaceous raw material. Hence, the present invention provides a process that allows the production of alcohols from a carbonaceous raw material, with improved carbon efficiency, superior selectivity and, in particular, with a more efficient conversion into ethanol.

Figure 1 represents an embodiment of a process scheme according to the present invention, wherein the references correspond to those used in the present description and the appended claims.

For the purpose of the present invention and the appended claims, the term 'approval' is used to describe the process by which an alcohol with n carbon atoms (eg methanol) enters the global process and an alcohol containing n + 1 carbon atoms (eg ethanol) exits the global procedure.

For the purposes of the present invention and the appended claims, "alkyl esters" are a class of chemical compounds and functional groups and consist of a carboxylic acid in which at least one -OH (hydroxy) group is replaced by a group - O-alkyl (alkoxy); The alkyl esters produced according to the present invention are preferably methyl ethanoate, ethyl ethanoate, methyl propanoate, ethyl propanoate, propyl butanoate and / or butyl pentanoate.

Thus, the present invention relates to a method for the approval of an alcohol or alcohols, characterized by the following steps:

one.

 introducing an alcohol or alcohols into an etherification unit to produce an ether or alkyl ethers,

2.

 introducing at least a portion of the ether or alkyl ethers from step 1, together with CO and optionally H2, into a carbonylation unit, in the presence of a homogeneous or heterogeneous acid carbonylation catalyst, to produce an alkyl ester or esters,

3.

 introducing at least a portion of the ester or alkyl esters from step 2, together with H2, into a hydrogenation unit, to produce an approved alcohol or alcohols,

Four.

optionally recycle at least a portion of the alcohol or approved alcohols from stage 3 to the etherification unit of stage 1, and

5.

 recovering an alcohol or alcohols from the hydrogenation unit of step 3.

Where the alcohol or alcohols introduced into the etherification unit (ie in step 1) can be produced by any appropriate method known to those skilled in the art, e.g. ex. through fermentation (described above herein).

For the purposes of the present invention and the appended claims, the following terms are defined hereinafter:

-

The 'dew point temperature' is a threshold temperature, for example, for a given pure component

or a mixture of components, at a given pressure, if the system temperature rises above the dew point temperature, the mixture will exist as a dry gas. Also, below the dew point temperature, the mixture will exist as a vapor that contains some liquid.

-

"Gas" and / or "gas phase" are defined as a pure component, or a mixture of components, that are above the dew point temperature.

-

’The hourly space gas velocity '(GHSV) is defined as the volume of gas fed per unit volume of catalyst per hour, at standard temperature (0 ° C) and pressure (0.101325 MPa).

-

’The liquid hourly space velocity’ (LHSV) is defined as the volume of liquid fed per unit volume of catalyst per hour.

The invention described above in the present invention has application for the homologation of existing raw materials derived from petroleum and biomass, namely alcohols, but can also be used to standardize ethers and esters. The usefulness of the process can be further extended by including a further processing step, for example carboxylic acids can be homologated by adding an esterification stage to convert them into a suitable ester intermediate. Similarly, olefins can be homologated by this procedure by adding a hydration stage to generate a matter

alcohol premium In particular, this invention finds application in the generation of fuels and improved petrochemical raw materials. For example, the approval of methanol by a step through the described process can give ethanol, a valuable combination raw material for gasoline. Additionally, an additional step (recycling of methanol and some ethanol) through the process described hereinbefore can generate n-propanol, a material with combination properties for gasoline (eg calorific value and vapor pressure) superior. The passage of n-propanol through the procedure described hereinbefore gives butanol or butanoles having desirable properties as combination components for gasoline and as intermediates in chemicals and fuels, e.g. ex. din-butyl ether has a high cetane number for combination. The starting alcohol, as described, need not only be methanol but may include other alcohols and mixtures such as bioethanol or other bioalcohols (eg methanol, ethanol, propanol and butanol). The use of bioalcohols in this procedure is advantageous since the approval will generally improve the combination properties of alcohols in fuels and thereby improve the properties of combined fuels and thus allow the incorporation of higher levels of renewable raw materials than possible else. This is advantageous in an environment where reduced greenhouse gas emissions and improved sustainability are important objectives. This process can also directly give valuable derivatives of alcohols such as dibutyl ether, ethylpentyl ether which could find application as a combination component for diesel oils, and valuable chemicals such as propyl propanoate and ethyl ethanoate (valuable solvents) and isobutanol, a precursor for isobutene.

Alcohol raw materials can be obtained from synthesis gas by methods previously described herein. These raw materials can also be obtained by hydrating olefins, in which case the higher olefins give secondary and tertiary alcohols, e.g. ex. Propene predominantly gives isopropanol and isobutene will give t-butanol. Typical sources of such olefin streams include catalytic cracking of hydrocarbons as commonly practiced in oil refineries, methanol to olefin (MTO) and oxygenated to olefin (OTO) processes, and dismutation reactions. Thus, including a hydration stage or an alcohol addition stage (in which case an ether raw material is obtained), these by-product streams can be usefully used.

The process is particularly advantageous for these raw materials since it produces, for C3 olefins and higher, desirable branched primary alcohol products not obtainable by direct hydration of the corresponding olefin. For example, propene can be converted into isobutanol and tertiary butanol (2-methylpropan-2-ol) into 2,2-dimethylpropan-1-ol.

The present invention also relates to a process for the conversion of a carbonaceous raw material into an alcohol or alcohols, in which the carbonaceous raw material is first converted into synthesis gas; which later becomes an alcohol or approved alcohols, characterized by the following stages:

one.

 introducing a carbonaceous raw material into a synthesis gas generation unit, in order to produce a mixture of carbon oxide or oxides and H2 (synthesis gas),

2.

 introducing at least a portion of the oxide or oxides of carbon and H2, from step 1, into an alcohol synthesis unit, to produce an alcohol or alcohols,

3.

 introducing at least a portion of the alcohol or alcohols from step 2 into an etherification unit to produce an ether or alkyl ethers,

Four.

 introducing at least a portion of the ether or alkyl ethers from step 3, together with CO and optionally H2, into a carbonylation unit, in the presence of a homogeneous or heterogeneous acid carbonylation catalyst, to produce an alkyl ester or esters,

5.

 introducing at least a portion of the ester or alkyl esters from step 4, together with H2, into a hydrogenation unit, to produce an approved alcohol or alcohols,

6.

 optionally recycle at least part of the alcohol or approved alcohols from stage 5 to the etherification unit of stage 3, and

7.

 recover the alcohol or alcohols sought from the hydrogenation unit of step 5.

According to one aspect of the present invention, the synthesis gas raw material, a mixture of oxide and oxides of carbon and H2, which is used to produce the feed stream of methanol, is preferably produced from a carbonaceous raw material.

The carbonaceous raw material is preferably a material such as biomass, plastic, naphtha, refinery glues, gas

of gross synthesis (from underground coal gasification or biomass gasification), exhaust gas from smelting furnaces, municipal waste, coal layer methane, coal and / or natural gas, with coal and natural gas being the preferred sources . For a person skilled in the art, a combination of sources, for example coal and natural gas, can also be used to advantageously increase the ratio of H2 to carbon.

Natural gas contains a range of hydrocarbons (eg C1-C3 alkanes), in which methane predominates. In addition to this, natural gas will usually contain nitrogen, CO2 and sulfur compounds. Preferably, the nitrogen content of the raw material is less than 40 mol%, more preferably less than 10 mol% and most preferably less than 2 mol%.

Procedures for producing synthesis gas in a synthesis gas plant are well known. Each method has its advantages and disadvantages, and the choice of using a particular reforming procedure over another is governed by economic considerations and the available feed currents, as well as by the desire to obtain the molar ratio (H2-CO2) :( Optimum CO + CO2) in the resulting synthesis gas that is suitable for additional chemical processing. An analysis of available synthesis gas production technologies is provided both in Hydrocarbon Processing, 1999, 78: 4, 87-90 and 92-93 and in Petrole et Techniques, 1998, 415, 86-93.

It is also known that synthesis gas can be obtained by catalytic partial oxidation of hydrocarbon material in a microstructured reactor as exemplified in IMRET 3: Proceedings of the Third International Conference on Microreaction Technology, ed. W. Ehrfeld, Springer Verlag, 1999, pages 187-196. Alternatively, the synthesis gas can be obtained by a partial catalytic oxidation of short contact time of hydrocarbon feedstocks as described in EP 0303438. The synthesis gas can also be obtained through a 'compact reformer' process as described described in Hydrocarbon Engineering, 2000, 5: 5, 67-69; Hydrocarbon Processing, 2000, 79: 9, 34; Today’s Refinery, 2000, 15: 8, 9; WO 9902254 and WO 0023689.

Typically, for the production of commercial synthesis gas, the pressure at which the synthesis gas is produced from a water vapor reformer ranges from about 0.1 to 10 MPa, preferably from 2 to 3 MPa and the temperatures at that the synthesis gas leaves the reformer vary from approximately 700 to 1000 ° C. Also, for the production of commercial synthesis gas, the pressure at which the synthesis gas is produced from an autothermal reformer varies from about 0.1 to 10 MPa, preferably from 2 to 5 MPa, and the temperatures at which the Synthesis gas leaves the reformer range from approximately 700 to 1300 ° C. When high temperatures are necessary in order to produce a favorable equilibrium for the production of synthesis gas, and to avoid metallurgy problems associated with carbon sprinkling. The synthesis gas contains a molar ratio of (H2-CO2) :( CO + CO2) that varies from 0.8 to 3.0, which depends on the carbonaceous raw material or materials and the reforming method used. For example, when natural gas is used as the carbonaceous raw material for steam reforming, the synthesis gas obtained usually has a maximum ratio of (H2-CO2) :( CO + CO2) of 3.0. However, when natural gas is used as the carbonaceous raw material for autothermal reforming, the synthesis gas obtained usually has a ratio of (H2-CO2) :( CO + CO2) of 1.5.

According to a preferred embodiment of the present invention, the molar ratio, (H2-CO2) :( CO + CO2), of the synthesis gas stream leaving the synthesis gas generating unit or units is greater than 1, 6, more preferably greater than 1.8 and most preferably greater than 2.0. Preferably, the molar ratio, (H2-CO2) :( CO + CO2), of said synthesis gas stream leaving the synthesis gas generating unit or units is less than 3.0, preferably less than 2, 75, more preferably less than 2.4 and most preferably less than 2.2.

Although the aim of the present invention is to homologize alcohols to produce higher alcohols, it is obvious to the person skilled in the art that the present invention may also allow searching for additional products, such as alkyl ethers and / or alkyl esters and / or carboxylic acids, together with said alcohols, wherein these additional products mentioned are present as intermediates and / or precursors in the process of the present invention. When the ester and / or acid products are sought, then the (H2-CO2) :( CO + CO2) of the required synthesis gas is preferably less than the above values.

According to another embodiment of this invention, when the carbonaceous raw material used for the generation of synthesis gas is not an aliphatic hydrocarbon (eg coal, aromatic material, biomass), the molar ratio (H2-CO2) :( CO + CO2) of the outlet synthesis gas is preferably adjusted to the desired value by the addition of H2 or the removal of CO2.

The CO2 can be removed by using a simple, yet efficient separation method, known to those skilled in the art, for example, a "membrane separation method." Such membrane technologies can be found in ’Purification and Recovery Options for Gasification’ D. J. Kubek, E. Polla, F. P. Wilcher, UOP, 1996.

Alternatively, the CO2 can be recovered and removed by any suitable method or methods known to those skilled in the art, for example, by reacting with amines; by washing with methanol (ie, the RECTISOL procedure) and / or using hot potassium carbonate (eg the BENFIELD procedure).

According to a preferred embodiment of the present invention, the output stream obtained from the synthesis gas reactor (eg using a water vapor reformer) essentially comprises a mixture of oxide and oxides of carbon and H2. It can also comprise water, nitrogen and traces of unconverted hydrocarbons (eg C1-C3 alkanes).

According to a preferred embodiment of the present invention, during the generation of synthesis gas, an additional stage can be employed whereby the raw material is first purified to remove sulfur and other potential catalyst poisons (such as halides or metals, eg mercury) before becoming a synthesis gas; alternatively, this treatment can also be carried out after the preparation of the synthesis gas, for example, when coal or biomass is used.

According to an embodiment of the present invention, at least part of said synthesis gas stream is then introduced into an alcohol synthesis unit, in order to produce a stream comprising alcohol or alcohols. Preferably, the molar ratio, (H2-CO2) :( CO + CO2), of said synthesis gas feed stream fed to the alcohol synthesis unit is greater than 1.6, more preferably greater than 1.8 and most preferably greater than 2.0. Preferably, the molar ratio, (H2-CO2) :( CO + CO2), of said synthesis gas feed stream fed to the alcohol synthesis unit is less than 3.0, more preferably less than 2.5 and most preferably less than 2.2.

When the alcohol synthesis unit is used to produce methanol, it is preferably operated at a temperature of more than 200 ° C, more preferably more than 220 ° C and most preferably more than 240 ° C; and preferably less than 310 ° C, more preferably less than 300 ° C and most preferably less than 290 ° C. Likewise, it is operated at a pressure preferably of more than 2 MPa and most preferably more than 5 MPa; and preferably less than 10 MPa and most preferably less than 9 MPa. In fact, since the synthesis of methanol is an exothermic reaction, the chosen operating temperature is governed by a weighting promoting the direct reaction (i.e. not adversely affecting the equilibrium) and helps the conversion rate (i.e. superior productivity).

The catalysts used for the synthesis of methanol in particular can be divided into two groups:

i. high pressure zinc catalysts, composed of zinc and a promoter; Y

ii. Low pressure copper catalysts, composed of zinc oxide, copper oxide and a promoter.

Hence, according to a preferred embodiment of the present invention, a preferred methanol synthesis catalyst is a mixture of copper, zinc oxide and a promoter such as chromia or alumina. Under the aforementioned operating conditions, these mentioned mixtures can catalyze the production of methanol from CO and H2 with high selectivity.

When the alcohol synthesis unit is used to produce a mixture of alcohols, such as methanol, ethanol, propanol or propanoles (n-propanol with low amounts of isopropanol) and butanol or butane (n-butanol and isobutanol), the conditions of Operation of the alcohol synthesis unit may vary depending on the catalyst employed. Suitable catalysts include modified methanol synthesis catalysts, modified Fischer-Tropsch catalysts, mixed metal oxide catalysts, rhodium based catalysts and molybdenum sulphide based catalysts.

According to a preferred embodiment of the present invention, the alcohol synthesis unit can be any reactor that is suitable for producing alcohol, for example a fixed bed reactor, which can be operated with or without external heat exchange equipment, e.g. . ex. a multitubular reactor; or a fluidized bed reactor; or a hollow reactor.

Molybdenum sulfide based catalysts are preferred; These can be modified by a promoter. The promoter or promoters can be added as salts during catalyst preparation; The preferred promoter or promoters are potassium ions and are derived from a potassium salt, such as potassium carbonate or ethanoate. Preferred charges of potassium ions per molybdenum are between 0.7 and 1.5, most preferably between 1.0 and 1.4.

The preferred catalyst for the synthesis of alcohols, according to the present invention, is a cobalt-containing molybdenum sulfide catalyst, the ratio of cobalt to molybdenum preferably being comprised

between 0.5 and 3.0, more preferably between 0.5 and 1.0 and most preferably between 0.5 and 0.9.

When catalysts based on molybdenum sulphide are used, the alcohol synthesis unit is preferably operated at a temperature of more than 150 ° C, preferably more than 250 ° C and most preferably more than 280 ° C; and less than 400 ° C, preferably less than 350 ° C and most preferably less than 320 ° C. Likewise, it is operated at a pressure of more than 2 MPa and preferably more than 8 MPa; and less than 20 MPa and preferably less than 15 MPa. According to a preferred embodiment of this aspect of the present invention, the current leaving the alcohol synthesis unit is subsequently purified, by any method known to those skilled in the art, to remove any by-products such as alkanes, aldehydes, ketones and / or water.

According to the present invention, the alcohols produced in the alcohol synthesis unit are then introduced into an etherification unit to produce an ether or alkyl ethers. For example, dimethyl ether is commonly produced by the etherification of methanol catalyzed by Lewis or Bronsted acids; numerous catalysts have been presented for this reaction, e.g. ex. Houben-Weilo, vol. VI / 3, part 3, Georg Thieme Verlag, Stuttgart 1965, pp. 17, 18. Dr. Alexander Wacker, Gesellschaft für elektrochemische Industrie GmbH, DE 680328, 1934 (P. Halbig, O. Moldenhauer). R. L. Brown, W. W. Odells, US 1873537, 1927. N. V. de Bataafsche Petroleum Maatschappij, FR 701335, 1930; GB 332756, 1929; GB 350010, 1931; GB 403402, 1932. Additional examples of etherification catalysts include: iron chloride; copper sulphate; copper chloride; manganese chloride; aluminum chloride; aluminum sulfate; chromium sulfate; alums; thorium compounds; aluminum oxide; titanium oxide; barium oxide; Silica gel; aluminum phosphate; and acidic ionic liquids. The etherification catalyst chosen can be either homogeneous or heterogeneous.

According to the present invention, preferred etherification catalysts are heterogeneous catalysts, such as aluminum oxides and aluminum silicate, which can be modified by impurification. Corresponding etherification catalysts and operating conditions of the process that can be advantageously used according to the present invention are described in Mobil Oil Corporation, DE 2818831, 1978 (F. G. Dwyer, A-. B. Schwartz); DE-OS 3201155, 1982 (W. K. Bell, C. Chang); Du Pont, EP-A 99676, 1983 (D. L. Brake); Mitsubishi Chemical Industries, EP-A 124078, 1984 (N. Murai, K. Nakamichi, M. Otake, T. Ushikubo).

Zeolites, strongly acidic ion exchange resins, supported heteropolyacids (such as silicovolframic acid) and mixtures thereof can also be advantageously used as etherification catalysts; and supported heteropolyacid catalysts are preferably used according to the present invention.

During the etherification process of alcohols to produce an ether or alkyl ethers, e.g. ex. methanol to dimethyl ether (DME) and / or methanol and ethanol to dimethyl ether and methyl ether and diethyl ether, water is also produced. It is preferred according to the present invention to proceed with the removal of said water before the introduction of the ether or alkyl ethers to the next stage of this invention. Other ethers can also be produced during the etherification process depending on the composition of the alcohol raw material introduced into the etherification unit, e.g. For example, dipropyl ether, dibutyl ether, dipentyl ether and ethylpentyl ether can also be produced and some of said ethers can advantageously be separated, isolated and used as additives for diesel oils.

According to the present invention, at least part of the ether or alkyl esters mentioned together with CO and optionally H2 is introduced into a carbonylation unit in the presence of a homogeneous or heterogeneous acid carbonylation catalyst. Applicants have found a preferred embodiment, for which it is especially advantageous (eg increased catalyst duration) to carry out the carbonylation process in the presence of H2. Preferably, at least part, most preferably all, of said ether or alkyl ethers emanate from the said etherification unit. However, according to an alternative embodiment of the present invention, at least part of the alkyl ether introduced into the carbonylation unit can also be originated from a direct process of synthesis gas to alkyl ether (eg synthesis gas up to dimethyl ether). For example, the reaction to produce dimethyl ether is exothermic and is controlled by equilibrium, like the methanol synthesis reaction; however, the reaction to produce dimethyl ether is more thermodynamically favorable than methanol synthesis, under methanol synthesis reaction conditions. This has led to the close coupling of these reactions in a single reactor as a means to increase the potential yield per step. This can be achieved by adding acidity to a methanol synthesis catalyst, or by mixing an acid catalyst with a methanol synthesis catalyst. Such procedures are often referred to as direct synthesis, e.g. ex. Snamprogetti, SpA., DE 2 362944, 1973 (G. Giorgio); DE 2757788, 1977 (G. Manara, B. Notari, V. Fattore); DE 3220547, 1982 (G. Manara). The same coupling can also be applied to the synthesis of mixed alcohols, for example, the inclusion of a subsequent acid bed in a mixed alcohol synthesis unit can be used to produce an ether or alkyl ethers.

Suitable carbonylation catalysts used in the present invention are homogeneous acid catalysts.

or heterogeneous acids. For example, the catalyst can be selected from strong acid catalysts.

homogeneous, p. ex. phosphoric acid; and heterogeneous analogs such as phosphoric acid catalysts and SAPO supported. Other heterogeneous catalysts include zeolites; strongly acidic ion resin catalysts; and mixtures thereof.

According to an embodiment of the present invention, the carbonylation process for producing a product comprising an alkyl ester or esters of a lower aliphatic carboxylic acid comprises reacting an ether.

or alkyl ethers with carbon monoxide under substantially anhydrous conditions in the presence of hydrogen and a catalyst comprising a zeolite having at least one 8-member ring channel, said channel of an 8-member ring being interconnected with a channel defined by a ring with more than or equal to 8 members, said 8-member ring having a gap size of at least 2.5 angstroms x at least 3.6 angstroms and at least one Brønsted acid site and where the zeolite has a silica ratio: X2O3 of at least 5, wherein X is selected from aluminum, boron, iron, gallium and mixtures thereof, such as mordenite or ferrierite.

A component of the carbonylation process feed comprises (mainly) an ether or alkyl ethers, which according to the present invention are produced during the above-mentioned etherification process. Said ether or alkyl ethers are compounds having the formula

R1-O-R2

wherein R1 and R2 are independently C1-C6 alkyl groups. The total number of carbon atoms in groups R1 and R2, if R1 and R2 are alkyl groups, is from 2 to 12, preferably from 2 to 8, most preferably from 2 to 6. Preferably, R1 and R2 are alkyl groups linear chain, most preferably linear chain alkyl groups having 1 to 3 carbon atoms each.

The global carbonylation reaction can be represented as

 R1-O-R2 + CO � R1COOR2

The term "alkyl", as used herein, means a saturated linear or branched chain aliphatic group, or a combination thereof, which has the indicated number of carbon atoms (ie C3 means three carbon atoms ). Examples of acyclic alkyl groups include groups such as methyl, ethyl, n-propyl, isopropyl, n-butyl, t-butyl, isobutyl, sec-butyl, and the various pentyl and hexyl isomers.

The term "alkylene", as used herein, refers to saturated aliphatic moieties that can form two single bonds with other moieties. This group includes, for example, methylene (-CH2-), ethylene (-CH2CH2-) and hexylene [(-CH2-) 6]. Although alkylene groups may be straight or branched chain groups, straight chain alkylene groups are preferred for use in the carbonylation process of this invention.

If the alkyl ether is a symmetric ether, e.g. ex. dimethyl ether, the main product will be the corresponding alkyl ester of an aliphatic acid (in this almost methyl ethanoate). If the ether is asymmetric, the product will comprise one

or both of the two possible carboxylic acid esters, depending on which of the two C-O bonds is cleaved in the reaction. For example, if the feed is methyl ethyl ether (R1 = methyl; R2 = ethyl), then the product will comprise ethyl ethanoate and / or methyl propanoate.

The carbonylation unit feed may comprise substantially pure carbon monoxide (CO), for example, carbon monoxide typically provided by industrial gas suppliers, or it may contain impurities that do not interfere with the conversion of the alkyl ether into the desired ester, such as nitrogen, helium, argon, methane and / or carbon dioxide. For example, the feed may comprise CO that is typically made commercially by removing hydrogen from the synthesis gas (produced during the above-mentioned synthesis gas generation stage), through a cryogenic separation and / or the use of a membrane.

In fact, according to a preferred embodiment of the present invention, the carbon monoxide feed to the carbonylation unit contains substantial amounts of hydrogen. For example, the feed introduced into the carbonylation unit may be what is commonly known as synthesis gas, and is preferably derived from the said synthesis gas generating unit.

As indicated, the carbonylation process takes place in the presence of both carbon monoxide and hydrogen; preferably, the ratio of carbon monoxide: hydrogen is in the range of 1: 3 to 15: 1 on a molar basis, such as 1: 1 to 10: 1.

Zeolites, both natural and synthetic, are microporous aluminosilicate materials that have a defined crystalline structure as determined by X-ray diffraction. The chemical composition of zeolites can vary widely, but typically consist of SiO2 in which some of the atoms of Si se

they can replace by tetravalent atoms such as Ti or Ge, by trivalent atoms such as Al, B, Ga, Fe or by bivalent atoms such as Be, or by a combination thereof. A zeolite is comprised of a channel system that can be interconnected with other channel systems or cavities such as side bags or cages. Channel systems are of uniform size within a specific zeolite and can be three-dimensional, but they are not necessarily so and can be two-dimensional or one-dimensional. Channel systems of a zeolite are typically accessed through 12-member rings, 10-member rings or 8-member rings. The specific zeolite mentioned hereinabove contains at least one channel that is defined by an 8-member ring. Preferred zeolites are those that do not have side bags or cages within the zeolite structure. The Atlas of Zeolite Framework Types (C. Baerlocher, WM Meier, DHOlson, 5th ed. Elsevier, Amsterdam, 2001) together with the Internet version (http://www.iza-structure.org/databases/) is a compendium of topological and structural details about the ordinations of the zeolites, including the types of annular structures present in the zeolite and the dimensions of the channels defined by each type of ring. For the purposes of the present invention, the term "zeolite" also includes materials having a zeolite-like structure such as de-stratified porous crystalline oxide materials such as ITQ-6 and columnar laminated oxide materials such as ITQ-36.

The carbonylation process of the present invention preferably employs a zeolite having at least one channel defined by an 8-membered ring of tetrahedrally coordinated atoms (tetrahedra) with a hole size having a minimum dimension of 2.5 angstroms x 3, 6 angstroms. The channel of the 8-member ring is interconnected with at least one channel defined by a ring with equal to or more than 8 members, such as 10 and / or 12 members. The interconnected 8, 10 and 12 member ring channels provide access to Brønsted acid sites contained in the 8 member ring channels to allow the carbonylation of the ether or alkyl ethers, such as dimethyl ether, to advance at acceptable rates.

The zeolite for use in the carbonylation process of the present invention may consist of interconnected channels defined only by 8-member rings, such as zeolites of the CHA sort type, for example, chabazite, and the ITE sort type, for example ITQ-3 Preferably, however, the zeolite has at least one channel formed by an 8 member ring and at least one interconnection channel defined by a ring with more than 8 members, such as a 10 and / or 12 member ring. Non-limiting examples of zeolites having 8-member ring channels and upper interconnecting ring channels include zeolites of a MOR sort type, for example, mordenite, FER, such as ferrierite and ITQ6, OFF, for example, offretite, GME , for example gmelinite, MFS, such as ZSM-57, EON such as ECR-1, and ETR such as ECR-34. Preferably, the zeolites for use in the process of the present invention have at least one channel of a ring of 8 members interconnected with at least one channel of a ring of 12 members, such as those of sort type MOR, OFF and GME , for example, mordenite, offretite and gmelinite.

The opening (pore width) of a channel of an 8-member ring of the zeolite has a minimum dimension of 2.5 x 3.6 angstroms. The dimensions of the channels of the zeolite sort types can be found, for example, in the Atlas of Zeolite Framework Types. In addition, M.D. Foster, 1. Rivin, M.M.J. Treacy and O. Delgado Friedrichs in "A geometric solution to the largest-free-sphere problem in zeolite frameworks" Microporous and Mesoporous Materials 90 (2006) 32-38, have used Delaunay triangulation methods applied to known zeolite arrays and have tabulated the larger diameters of the free spheres for diffusion along the three main crystallographic directions for the ordinations of zeolite 165 that are currently listed in the Atlas of Zeolite Framework Types. Ring gap sizes can be modified by suitable atomic substitutions that change both the bond lengths and the bond angles of the tetrahedrally coordinated atoms and the bridge oxygen.

The following is a partial list of types of zeolite management that have at least one interconnected 8-member ring channel of a minimum dimension of 2.5 x 3.6 angstroms taken from The Atlas of Zeolite Framework Types:

MOR

Mordenite 12 (6.5 x 7.0 Å) 8 (3.4 x 4.8 Å) 8 (2.6 x 5.7 Å)

OFF

Offretita 12 (6.7 x 6.8 Å) 8 (3.6 x 4.9 Å)

FER

Ferrierite 10 (4.2 x 5.4 Å) 8 (3.5 x 4.8 Å)

CHA

Chabazita 8 (3.8 x 3.8 Å)

ITE

ITQ3 8 (3.8 x 4.3 Å) 8 (2.7 x 5.8 Å)

GME

Gmelinite 12 (7.0 x 7.0 Å) 8 (3.6 x 3.9 Å)

ETR

ECR-34 18 (10.1 Å) 8 (2.5, x 6.0 Å)

MFS

ZSM-57 10 (5.1 x 5.4 Å) 8 (3.3 x 4.8 Å)

EON

ECR-1 12 (6.7 x 6.8 Å) 8 (3.4 x 4.9 Å) 8 (2.9 x 2.9 Å)

Zeolites are available from commercial sources. Alternatively, they can be synthesized using known techniques. In general, synthetic zeolites are prepared from aqueous reaction mixtures comprising appropriate sources of oxides. Organic director agents may also be included in the reaction mixture for the purpose of influencing the production of a zeolite having the desired structure. After the components of the reaction mixture are properly mixed with each other, the reaction mixture is subjected to appropriate crystallization conditions. After crystallization of the reaction mixture is complete, the crystalline product can be recovered from the rest of the reaction mixture. Such recovery may involve filtering the crystals, washing with water, followed by a high temperature calcination treatment. The synthesis of zeolites is described in numerous references. For example, zeolite Y and its synthesis is described in US 3,130,007, zeolite ZSM-23 is described in US 4,076,842 and J.Phys. Chem. B, 109, 652-661 (2005), Zones, S.I. Darton, R.J., Morris, R and Hwany, S-J; ECR-18 is described in Microporous Mesoporous Mat., 28, 233-239. (1999), Vaughan D.E.W. & Strohmaier, K.G .; Theta-1 is described in Nature, 312, 533-534 (1984). Barri, S.A.I., Smith W.G., White, D and Young, D .; mazzita is described in Microporous Mesoporous Mat., 63, 33-42 (2003), Martucci, A, Alberti, A, Guzmar-Castillo, M.D., Di Renzo, F and Fajula, F .; Zeolite L is described in Microporous Mesoporous Mat., 76; 81-99 (2004), Bhat, S.D., Niphadkair, P.S., Gaydharker, T.R., Awate, S.V., Belhekar, A.A. and Joshi, PN and also in J. Ind. Eng. Chem. Vol. 10, No. 4 (2004), 636-644, Ko YS, Ahn WS and the offretite is described in Zeolites 255-264, Vol. 7, 1987 Howden MG

The zeolite catalyst for use in the carbonylation process of the present invention is used in acidic form, generally referred to as the 'H' form of the zeolite, for example, mordenite H, ferrite H. Other forms of the zeolite, such as the NH4 form can be converted into the H form, for example, by calcining the NH4 form at elevated temperature. The acid form of a zeolite will possess Brønsted acid sites (H +) that are distributed among the various channel systems in the zeolite. For example, mordenite H has H + sites located in the 12-member ring channels and in the 8-member ring channels. The number or concentration of H + species residing in any particular channel system can be determined by known techniques such as infrared and NMR spectroscopic techniques. The quantification of Brønsted acidity by FTIR and NMR spectroscopy is described, for example, in Makarova, MA, Wilson, AE, van Liemt, BJ, Mesters, C. de Winter, AW, Williams, C. Journal of Catalysis 1997, 172, (1), 170. The two types of channels in mordenite H (defined by 12-member rings and 8-member rings) give rise to at least two bands associated with the hydroxyl region of mordenite H, a corresponding to the vibration in the larger pores and the other, at a lower frequency, which vibrates in the smaller pores. It has been observed that there is a correlation between the number of H + sites located in a channel of an 8-member ring and the rate of carbonylation while no such correlation has been observed for 12-member ring channels. It has been found that carbonylation rates increase in parallel with the number of H + sites within the 8-member ring channels. In contrast, a correlation with the number of H + sites within the 12-member ring channels is not evident. The number of H + sites within the 8-member ring channels can be controlled by replacing H + with metal cations such as Na + or Co2 + using known ion exchange techniques.

It can be expressed that the chemical composition of a zeolite (which can be used to catalyze the carbonylation process) implies the molar ratio:

SiO2: X2O3

wherein X is a trivalent element, such as aluminum, boron, iron and / or gallium, preferably aluminum. The SiO2: X2O3 ratio of a given zeolite is often variable. For example, it is known that mordenite can be synthesized with SiO2: Al2O3 ratios of 6 to 90 or more, zeolite Y, from about 1 to about 6, chabazite of about 2 to 2,000 and gmelinite can be synthesized with SiO2: Al2O3 ratios of more than 4. In general, the upper limit of the SiO2: X2O3 ratio is not limited, for example, the zeolite ZSM-5. The specific zeolite mentioned hereinbefore has a SiO2: X2O3 molar ratio of at least 5, preferably in the range of 7 to 40, such as 10 to 30. Suitably, the SiO2: X2O3 molar ratio is less than or equal to 100. Particular SiO2: X2O3 ratios can be obtained for many zeolites by desalumination (where X is Al), by standard techniques using high temperature current treatment

or acid wash.

In one embodiment, the catalyst used in the carbonylation unit is composed of mordenite or ferrierite, or mixtures or combinations of the two, either as such (ie, in acidic form, generally referred to as mordenite H and ferrierite H), or optionally ionically exchanged or charged with one or more metals such as copper, nickel, iridium, rhodium, platinum, palladium or cobalt. Mordenite catalysts may contain, in addition to silicon and aluminum atoms, additional elements in the arrangement of zeolite, particularly gallium and / or iron. Ferrierite catalysts may contain, in addition to silicon and aluminum atoms, additional elements in the arrangement of zeolite, particularly boron, gallium and / or iron. Sorting modifying elements for both types of catalysts can be introduced into the sorting by any conventional means. When a sort modifying element is used in a mordenite or ferrierite catalyst, the catalyst suitably has a ratio of the silica to the oxide of the sort modifying element of about 10: 1 to about 100: 1. The incorporation of T atoms where T is B, Ga or Fe in zeolites of the ferrite structure is disclosed in Melian-Cabrera et al., Catalysis Today 110 (2005) 255-263; Shawki et al., EP (Application) 234,766 (1987), Sulikowski et al., J. Chem. Soc., Chem. Comm., 1289 (1989); Borade et al., J. Chem Soc., Chem. Comm., 2267 (1996); Jacob et al., Zeolites 430 (1993) Vol. 13. The incorporation of T atoms into zeolites of the mordenite structure where the T atom is Ga or Fe is disclosed in Smith, WO 05/085162.

Mordenite (commonly available as mordenite Na, mordenite NH4 or mordenite H) is a member of the aluminosilicate zeolite mineral class. The mordenite formula in its Na form is usually given as Na (AlSi5O12) .3H2O or (Na2, Ca, K2) Al2Si10O24.7H2O. It is available from a number of commercial sources of material materials. Mordenite H has elliptical channels of 6.5 x 7.0 Å (12-member rings with hollow openings that go in the crystallographic direction c). It also has a system of smaller channels that go perpendicular to the 12-member ring channels (and that go in the b direction). These small channels consist of 3.4 x 4.8 Å channels that have 8-member ring holes of these dimensions. The mordenite structure also has a zigzag Y-branch of the pore structure due to the twisted 8-member rings (in the crystallographic direction c). This results in a distorted gap for each side of the Y branch of 2.6 x 5.7 Å. Ferrierite is another member of the aluminosilicate zeolite mineral class, also available in Na, NH4 and H forms. In the Na form its formula is generally given as Na0.8K0.2MgSi15Al3O or (Mg, Na2, K2, Ca ) 3-5Mg [Al5-7Si27.5-31O72] .18H2O. In addition, it is available from various commercial sources. Additional information on these and other materials can be found on the International Zeolite Association's website, www.iza-online.org. Ferrierite H has 10-member ring channels of 4.2 x 5.4 that go in the crystallographic direction c and 8-member ring channels of 3.5 x 4.8 that go in the direction b.

The carbonylation reaction must be carried out substantially in the absence of water when the specific zeolite mentioned above is used herein, for example by drying the zeolite catalyst before the start of the operation, for example, by preheating to 400-500 ° C .

In general, when a zeolite catalyst is used, the carbonylation process is carried out at temperatures at or below about 600 ° C, that is, at temperatures of about 150 to about 600 ° C, preferably about 250 at around 400 ° C. A feature of the process is that, surprisingly, the carbonylation of dimethyl ether (DME) to methyl ethanoate using catalysts based on mordenite zeolite and in substantial absence of water can be performed with very high selectivities.

Typical total operating pressures when zeolites are employed in the carbonylation process are from 1 MPa to about 20 MPa, preferably from about 1 MPa to about 15 MPa and most preferably from about 2.5 MPa to about 10 MPa; preferably with carbon monoxide pressures greater than 1 MPa and dimethyl ether pressures below 0.5 MPa

The carbonylation process with zeolites is carried out under substantially anhydrous conditions, that is to say in the substantial absence of water. The carbonylation of dimethyl ether to methyl ethanoate does not generate water in situ. Water has been found to inhibit the carbonylation of dimethyl ether to form methyl ethanoate. This is in comparison with prior art procedures, in which dimethyl ether was a coalition, and in which water was also fed to the reaction. Thus, water is kept as low as feasible, in order to allow

Let the desired reaction progress better. To achieve this, preferably, the alkyl ether and carbon monoxide reactants and the catalyst are dried before introduction into the process.

The carbonylation process can be carried out either as a continuous or discontinuous process, with continuous procedures being typically preferred. Essentially, the heterogeneous process is a gas phase operation, the reactants being introduced in either a liquid or a gas phase, and the products being extracted as gases. As desired, the reaction products can subsequently be cooled and condensed. The heterogeneous catalyst can be used as appropriate, either in a fixed bed or in a fluidized bed. By implementing the procedure, unreacted starting materials can be recovered and recycled to the unit. Next, the alkyl esters obtained as a product are fed to the hydrogenation unit; however, some alkyl esters (eg methyl ethanoate) can also be recovered and sold as such, or can be sent to other chemical processing units, as desired. Preferably, the entire reaction product can be sent to the hydrogenation unit for the conversion of the alkyl esters into the corresponding esters.

The alkyl ester product can be removed from the carbonylation unit by extracting the liquid reaction composition and separating the alkyl ester product by one or more stages of rapid and / or fractional development distillation of the other components. The alkyl ester product can also be removed as a vapor from the stream leaving the carbonylation unit. According to a preferred embodiment of the present invention, during the above separation stage, water is selectively removed from the carbonylation process in order to maintain the aforementioned operating conditions.

According to a preferred embodiment of the present invention, the ester or alkyl esters separated from the stream leaving the carbonylation unit are then purified to remove carboxylic acids (eg ethanoic acid) and water, before being introduced into the hydrogenation unit.

After purification and before introduction into the hydrogenation unit, the alkyl ester stream preferably contains less than 5% by weight of carboxylic acids, more preferably less than 1% by weight, even more preferably less than 0.1% by weight and most preferably less than 100 ppm by weight of carboxylic acids.

After purification and before introduction into the hydrogenation unit, the alkyl ester stream preferably contains less than 20 mol% water, preferably less than 2 mol%, more preferably less than 0.2 mol% and most preferably the hydrogenation unit is operated in the absence of water.

According to an alternative embodiment of the present invention, water represents between 0.5 and 20 mol%, more preferably between 0.5 and 15 mol% and most preferably between 1 and 5 mol% of the total liquid feed ( alkyl ester, alcohol and water) to the hydrogenation unit.

These mentioned purification treatments are especially advantageous when a copper based catalyst is used in the hydrogenation unit.

On the other hand, the alkyl ester feed is preferably vaporized before coming into contact with the hydrogenation catalyst, either by heating the alkyl ester in a separate vaporizer before it has contact with H2 or by heating the alkyl ester together with H2 (eg either in a separate container or in a bed prior to the hydrogenation unit). The feed mixture that includes recyclates that enter the hydrogenation unit (eg H2 together with the alkyl esters) is preferably introduced at more than 10 ° C above its dew point temperature, and more preferably at more than 20 ° C above its dew point temperature.

According to the present invention, the alkyl esters are introduced into the hydrogenation unit together with H2 to produce a stream comprising alcohols. In addition to the production of alcohols, the hydrogenation process also produces other reaction products such as alkyl esters arising from transesterification, trace amounts of methane, ethane, water and aldehydes, together with unreacted starting materials, e.g. ex. unreacted alkyl esters (eg from recycling) and unreacted H2.

For example, when methyl ethanoate is used as a feed for the hydrogenation unit, it also produces ethyl ethanoate by transesterification. The proportion of alkyl esters or esters present in the output stream will be determined by the nature of the catalyst and the degree of conversion. The proportion of alkyl esters or esters can be varied more, if desired, by introducing an acidic function into the catalyst bed to promote transesterification in situ. The ester or alkyl esters can be recovered from the hydrogenation unit for recycling to other processing stages such as the etherification unit, the carbonylation unit and / or as a recycle to the hydrogenation unit. More preferably, the amounts of

These flows are minimized.

According to a preferred embodiment of the present invention, in order to minimize the cost of separation and recycling of the ester or alkyl esters within the process, the hydrogenation unit is operated at a high conversion of the ester or alkyl ester feed. in an alcohol or alcohols, such as more than 50%, more preferably more than 75%, even more preferably more than 90% and most preferably more than 95%.

It has also been found that it is an advantage of this process that the selectivity of the hydrogenation process of the ester or alkyl esters in an alcohol or alcohols can be further increased, at the expense of undesirable by-products, such as the aforementioned alkanes (e.g. ethane and methane).

Preferably, at least a portion, preferably all, of the H2 fed to the hydrogenation unit emanates from the synthesis gas generation process (i.e., is obtained during the above-mentioned CO / H2 separation), where, if necessary, The H2 content can be further increased by subjecting said synthesis gas to a water-gas displacement reaction and a subsequent H2 separation.

Alternatively, the H2 stream can originate from a variety of other chemical processes, including ethereal crackers, styrene fabrication and catalytic reforming. However, it is known that the main commercial procedures for the intentional generation of an autothermal reforming of H are autothermal reforming, steam reforming and partial oxidation of hydrocarbon feedstocks such as natural gas, coal, coke, glues of desasfaltador, refinery and biomass waste. H2 can also be produced by water electrolysis.

The global choice of technology to produce H2 is generally determined by the economic considerations and the following factors:

i. raw material cost

ii. availability of raw materials

iii. capital cost

iv.

 local energy costs and operation; Y

v.

environmental considerations

The catalyst or catalysts used in the hydrogenation unit are selected from any of the following:

(i)

 a precious metal based catalyst, comprising at least one Group VIII noble metal of the periodic table (CAS version, for example iron, ruthenium, osmium, cobalt, rhodium, indium, nickel, palladium, platinum) and at least one of the metals chosen from rhenium, tungsten and / or molybdenum; and optionally an additional metal; which is capable of alloying with said noble metal of Group VIII; or

(ii)

 a copper-based catalyst (for example a copper chromite or a mixed catalyst based on copper-metal oxide in which the second metal can be copper, zinc, zirconium or manganese).

According to a preferred embodiment of the present invention, the catalyst or catalysts employed in the hydrogenation unit are a copper based catalyst, most preferably comprising copper and zinc. The copper-based catalyst comprising copper and zinc has a composition of copper oxide plus zinc oxide preferably greater than 75% by weight, more preferably greater than 90% by weight and most preferably greater than 95% by weight.

All of the above-mentioned catalysts can be advantageously supported on any support known to those skilled in the art; Non-limiting examples of such supports include carbon, silica, titania, clay, aluminas, zinc oxide, zirconia and mixed oxides. Preferably, the palladium-based catalyst is supported on carbon. Preferably, the copper-based catalyst is supported on zinc oxide and preferably comprises between 20 and 40% by weight of copper.

According to a preferred embodiment of the present invention, the catalyst or catalysts employed are heterogeneous.

The hydrogenation process can be carried out in a gas phase or phase regime

Soda / mixed liquid. The mixed gas / liquid phase regime is when the mixture of reactants, at the reaction conditions, is below the dew point temperature.

Hydrogenation can be carried out semi-continuously or continuously. The continuous operation mode is the most preferred.

The hydrogenation reaction can be carried out adiabically or isothermally; where the adiabatic mode is preferred. Suitable reactors include a single, or a plurality of, adiabatic bed reactor that can be used in series or parallel. For reactors used in series, heat exchangers and / or intercoolers and / or an additional reactant and / or recycled intermediate products between successive reactors can be used to control the reaction temperature. The preferred adiabatic temperature increase is less than 50 ° C, preferably less than 25 ° C and most preferably less than 10 ° C. The preferred use of adiabatic reactors is in series. Adiabatic reactors can be operated at different temperatures depending on the composition of the individual feeds to the reactors.

Hydrogenation can also be carried out in multitubular reactors, in which case a cooling / heating medium is circulated through the tubes to control the temperature. For exothermic reactions, as such, there will be a radial temperature gradient in the reactor, the preferred gradient is less than 50 ° C, preferably less than 25 ° C, most preferably less than 10 ° C. The preferred flow rate in this type of reactor is turbulent instead of laminar, this corresponds to a Reynolds number greater than 2,100 (where the velocity is approximated by the velocity in an unfilled tube).

The hydrogenation reaction can also be carried out in other types of reactors such as fluidized bed, rotating basket and Buss Loop heat exchange reactors. A mixed liquid / gaseous hydrogenation reaction can be carried out with coflux or counterflow of H2 and gas to the liquid (eg a bubbling reactor). The preferred mode of operation of gas / liquid reactors is the coflow, also known as percolator bed operation; This can be done in at least one tubular and / or multitubular reactor in series. The hydrogenation reaction may change from a mixed gas / liquid phase to a completely gas phase reaction, as the reaction descends through the reactor. Mixed phase hydrogenation can also be carried out in other types of reactors, or within a combination of different reactors, for example in a stirred or reservoir reactor with, or without, external circulation and optionally functioning as a cascade or reservoirs. stirred, a cyclic reactor or a mixer-reactor from Sulzer.

The hydrogenation reactor or reactors preferably operate at a temperature of more than 150 ° C, but less than 290 ° C.

According to a preferred embodiment of the present invention, the reaction temperature is greater than 150 ° C, preferably greater than 170 ° C and most preferably greater than 190 ° C; and less than 250 ° C.

The hydrogenation reaction can be operated at a pressure of more than 3 MPa, preferably at a pressure of more than 5 MPa; and at a pressure of less than 15 MPa, more preferably at a pressure of less than 13 MPa and most preferably at a pressure of less than 9 MPa.

According to an embodiment of the present invention, when the hydrogenation unit or units are operated with a copper-based catalyst, the feed mixture introduced into the reactor or reactors is always above its dew point temperature.

The GHSV for continuous operation may be in the range of 50 to 50,000 h-1, preferably 1,000 to

30,000 h-1 and most preferably 2,000 to 9,000 h-1.

The liquid ester substrate introduced into the hydrogenation unit preferably has an LHSV of less than 10 h-1, more preferably less than 5 h-1 and most preferably less than 3 h-1; For example, a typical LHSV for normal operation is approximately 1 h-1.

For the purpose of calculating the GHSV and / or the LHSV of the present invention, when the reactors are arranged in series, said "catalyst volume" refers to the combined volume of the catalyst in all reactors in series.

According to a preferred embodiment of the present invention, the molar ratio of H2 to alkyl ester that is introduced into the hydrogenation unit is greater than 2: 1, preferably the molar ratio is greater than 4: 1 and most preferably the molar ratio is greater than 5: 1; and is less than 100: 1, preferably less than 50: 1 and most preferably less than 15: 1.

According to a preferred embodiment of the present invention, the effluent leaving the hydrogenation reactor, (or the

last in a chain of reactors), is subjected to a separation stage, e.g. ex. by cooling the effluent and then feeding the cooled effluent to an instant vaporization vessel (in which said effluent can be cooled and / or vaporized instantly several times). After this treatment, a vapor stream containing the majority of H2 and inert gases (including light alkanes; nitrogen; CO; and CO2) in the effluent will be

5 recover for recycling to the unit or hydrogenation units. The concentration of said inert gases can be controlled by diverting a proportion of the gas recycle stream to a purge.

According to the present invention, the stream leaving the hydrogenation unit is then subjected to a separation stage (eg distillation), whereby a fraction comprising an alcohol or alcohols is separated and recovered. Said separated and recovered alcohols comprise the following: methanol, ethanol, propanol or

10 propanoles (n-propanol with low amounts of isopropanol), butanol or butane (n-butanol and isobutanol) and pentanol or pentaols and together represent the alcohols sought (as mentioned hereinbefore) of the present invention.

It should be noted that although all said temperature and pressure operating conditions form preferred embodiments of the present invention, they are by no means intended to be limiting, and the present invention

15 includes any other pressure and temperature conditions that achieve the same effect.

Claims (19)

1. A procedure for the approval of an alcohol or alcohols, characterized by the following stages:

one.

 introducing an alcohol or alcohols into an etherification unit to produce an ether or alkyl ethers,

2.

 introducing at least a portion of the ether or alkyl ethers from step 1, together with CO and optionally H2, into a carbonylation unit, in the presence of a homogeneous or heterogeneous acid carbonylation catalyst, to produce an alkyl ester or esters,

3.

 introducing at least a portion of the ester or alkyl esters from step 2, together with H2, into a hydrogenation unit, to produce an approved alcohol or alcohols,

Four.

optionally recycle at least a portion of the alcohol or approved alcohols from stage 3 to the etherification unit of stage 1, and

5.

 recovering an alcohol or alcohols from the hydrogenation unit of step 3.

2.

 A process according to claim 1, wherein the alcohol or alcohols introduced into the etherification unit in step 1 of claim 1 are prepared by the following steps:

one.

 introducing a carbonaceous raw material into a synthesis gas generation unit, in order to produce a mixture of carbon oxide or oxides and H2 (synthesis gas), and

2.

 introducing at least a portion of the oxide or the oxides of carbon and H2, from step 1, into an alcohol synthesis unit, to produce an alcohol or alcohols.

3.

 A process according to any one of the preceding claims, wherein the alcohol introduced into the etherification unit is methanol.

Four.

 A process according to any one of the preceding claims, wherein the etherification unit employs a catalyst or catalysts selected from the following: zeolites; strongly acidic ion exchange resins; supported heteropolyacids; or mixtures thereof.

5.

A process according to any one of the preceding claims, wherein the carbonylation process is carried out in the presence of a homogeneous acidic or heterogeneous acid catalyst, and is selected from a strongly homogeneous acidic catalyst or catalysts, e.g. ex. phosphoric acid; heterogeneous analogs - such as phosphoric acid catalysts and SAPO supported; and heterogeneous catalysts that include zeolites, strongly acidic ion resin catalysts; or a mixture thereof.

6.

A process according to any one of the preceding claims, wherein the carbonylation process is performed in the presence of a catalyst comprising a zeolite having at least one channel of an 8-membered ring.

7.

 A process according to claim 6, wherein the zeolite catalyst for use in the carbonylation process is used in acid form.

8.

 A process according to claim 6 or claim 7, wherein the carbonylation catalyst is composed of mordenite or ferrierite, or mixtures or combinations of the two.

9.

A process according to any one of the preceding claims, wherein the carbonylation process is performed under substantially anhydrous conditions.

10.

 A process according to any one of the preceding claims, wherein the carbonylation process is carried out at temperatures of 150 to 600 ° C.

eleven.

 A process according to any one of the preceding claims, wherein the hydrogenation unit employs a copper based catalyst.

12.

 A process according to any one of the preceding claims, wherein the alkyl ester stream entering the hydrogenation reactor contains less than 5% by weight of carboxylic acids.

13.

 A process according to any one of the preceding claims, wherein the ester current

alkyl that enters the hydrogenation unit contains less than 20 mol% of water.

14. A process according to any one of the preceding claims, wherein the hydrogenation is operated at a temperature that is greater than 150 ° C and less than 250 ° C.

 unit of

5

15. A process according to any one of the preceding claims, wherein the hydrogenation is operated at a pressure of more than 1 MPa and at a pressure of less than 15 MPa.  unit of